Hollow Fiber Membrane Module for Direct Contact Membrane Distillation-Based Desalination

ABSTRACT

Exemplary embodiments in desalination by direct contact membrane distillation present a cylindrical cross-flow module containing high-flux composite hydrophobic hollow fiber membranes. The present embodiments are directed to a model that has been developed to describe the observed water production rates of such devices in multiple brine feed introduction configurations. The model describes the observed water vapor production rates for different feed brine temperatures at various feed brine flow rates. The model flux predictions have been explored over a range of hollow fiber lengths to compare the present results with those obtained earlier from rectangular modules which had significantly shorter hollow fibers.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application is a divisional of U.S. non-provisional applicationSer. No. 16/607,015, filed on Oct. 21, 2019 and entitled “Hollow FiberMembrane Module for Direct Contact Membrane Distillation-BasedDesalination,” which is a national phase of international applicationserial number PCT/US2018/028860, filed on Apr. 23, 2018, and entitled“Hollow Fiber Membrane Module for Direct Contact MembraneDistillation-Based Desalination,” which itself claims priority to U.S.provisional application Ser. No. 62/488,437, filed on Apr. 21, 2017, andentitled “Hollow Fiber Membrane Module for Direct Contact MembraneDistillation-Based Desalination,” the contents of each of which beingincorporated herein in their entireties.

STATEMENT REGARDING FEDERALLY SPONSORED RESEARCH

This invention was made with government support under Agreement No.R12AC80907 awarded by the U.S. Bureau of Reclamation. The government hascertain rights in the invention.

FIELD OF THE DISCLOSURE

The present disclosure relates to a membrane module for direct contactmembrane distillation (DCMD). More particularly, the present disclosureis directed to a DCMD system and method for use in desalination.

BACKGROUND

Membrane distillation could be used for desalination, which is theproduction of fresh water from saline water. There are many potentialadvantages of membrane distillation for water production by suchdesalination techniques. These advantages include:

(a) membrane distillation produces high quality distillate;

(b) water can be distilled at relatively low temperatures (e.g., 30 to100 degrees C.) and low pressure (e.g., 1 atm);

(c) low grade heat (e.g., solar, industrial waste heat, or desalinationwaste heat) may be used; and

(d) water does not require extensive pretreatment to prevent membranefouling as in pressure-based membrane processes.

Generally, membrane distillation (MD) is an evaporation process of avolatile solvent or solute species from a solution (in most cases, anaqueous solution), driven by a difference between its partial pressureover the solution contacting one side of a porous hydrophobic membraneand its partial pressure on the other side of the membrane. When thepartial pressure difference through the membrane is created by thedirect contacting of a liquid cooler than the feed on the other side ofthe membrane, the process is called direct contact membrane distillation(DCMD). In a hollow fiber-based process, the hot brine flows on theshell side of the fiber and the cold distillate flows on the tube sidethrough the fiber bore.

In a MD process, the membrane is generally porous and hydrophobic. Inone variety of MD, direct contact membrane distillation (DCMD), hotbrine flows on one side of a gas-filled porous hydrophobic hollow fibermembrane and cold distillate flows on the other side of the membrane.Surface tension forces withhold liquids from the pores and preventpenetration by the liquids. The withholding of the liquids is intendedto prevent contact between the two liquids in a DCMD process. Generally,the solutions being processed are aqueous and their surface tensions arehigher than the critical surface tension of the polymeric membrane.

In a DCMD process, the temperature difference translates to acorresponding vapor pressure difference across the membrane and providesa driving force for the membrane distillation process. Evaporationoccurs at the solution surface if the vapor pressure on the solutionside is greater than the vapor pressure at the condensate surface.Vapors then diffuse through the pores to the cooler surface where theycondense.

Desalination by membrane distillation (MD) recovers pure water vaporfrom hot brine by passing the brine on one side of a porous hydrophobicmembrane whose pores are gas-filled. Direct contact MD (DCMD), vacuum MD(VIVID), sweep gas MD (SGMD) and air gap MD (AGMD) (Sirkar, 1992 [1];Lawson and Lloyd, 1997 [2]; Drioli et al., 2005 [3]; Khayet, 2008[4])are four different types of MD depending on the conditions maintained onthe other side of the membrane. In DCMD, cold distillate on the otherside of the membrane locally condenses water vapor coming through themembrane pores and becomes heated up in the process. This heat isrecycled to heat the cooled/spent brine from the DCMD unit in a heatexchanger for further desalination (Lee et al., 2011) [5].

Considerable research has been carried out on the DCMD process. A fewpublications are listed here: Schofield et al., 1987[6]; Schofield etal., 1990[7]; Martinez and Florido-Diaz, 2001[8]; Phattaranawik et al.,2003[9]; Alklaibi and Lior, 2006[10]; Khayet, 2008[4]. Extensive workhas also been carried out to characterize and scale up the DCMD process(Li and Sirkar, 2004[11]; Song et al., 2007[12]; Song et al., 2008[13])using a novel porous fluorosiloxane-coated porous hydrophobic hollowfiber membrane (HFM) housed in a rectangular cross-flow module as wellas develop its cost estimates (Gilron et al., 2007[14]) fordesalination. The water fluxes achieved were quite high. The membraneperformance was stable (Song et al., 2008) [13]. This last referencedescribes highly encouraging results from pilot plant studies forseawater desalination at a water production rate of around 2.34 L/min(0.62 gpm) using larger membrane modules; the salt concentrationachieved after continuous desalination for a few days using batchrecirculation was around 19%.

Extensive lab-scale studies with brines having highly supersaturatedsolutions of CaSO₄ and/or CaCO₃ have demonstrated excellent scalingresistance (He et al., 2008[15]; He et al., 2009a [16]; He et al., 2009b[17]) in these porous fluorosiloxane-coated polypropylene hollowfiber-based rectangular cross-flow DCMD modules and a countercurrentcascade of modules (Lee et al., 2011[5]). Brines having anti-scalants,e.g., reverse osmosis (RO) concentrates, did not lead to membranewetting (He et al., 2009b) [17].

The economics of desalination by a countercurrent cascade of cross flowmembrane distillation modules are strongly influenced by cascade designand energy cost. For an energy-efficient process with a countercurrentcascade of such cross-flow modules, the temperature difference betweenthe hot and cold streams in the countercurrent arrangement is low; thelarger the number of cross-flow stages, the lower the temperaturedifference. Using a countercurrent cascade, Lee et al. (2011) [5] haveexperimentally obtained a stage thermal efficiency value approaching90%. This reduction in conductive heat loss has in effect eliminated oneimportant shortcoming of the DCMD vis-à-vis VMD.

With low-cost steam, the technology appears to be near-competitive withreverse osmosis (Gilron et al., 2007[14]). Further, it can be used toconcentrate brine to around 20% salt (Song et al., 2008[13]) and therebyreduce the brine disposal cost in inland desalination. If waste heat orsolar heat sources are available, then the economics are even better.Produced water is an important energy source (Mondal and Wickramasinghe,2008[18]; Alkhudhiri et al., 2013[19]; Webb et al., 2009[20]). Such hotproduced waters are de-oiled first by induced gas/dissolved airflotation and walnut shell treatment. The de-oiled produced water,currently treated by as many as eight steps (in the OPUS™ Process)including substantial cooling, is treated by RO at the end (see Webb etal. (2009) [20] for a detailed process configuration). Using de-oiledproduced water (from Chevron Corp.), recent studies (Singh et al., 2013)[21] using small lab-scale modules achieved as much as 80% waterrecovery in one step via simple DCMD in batch recirculation mode; nocooling of the produced water was needed. Scaling problems werenonexistent. However, the current rectangular hollow fiber module designcould create a bottleneck for scaling up.

The rectangular membrane module structure has been described in Li andSirkar (2004) [11], Song et al. (2007) [12] and Song et al. (2008) [13].For example, in FIG. 3 of Li and Sirkar (2004) [11], part c shows anassembled module incorporating a picture frame containing the hollowfiber membranes in the middle. On each side of this picture frame, thereare two separate plastic pieces: the face box and the face plate. Thisdesign allowed hot brine to enter almost uniformly across the crosssection of the picture frame of HFMs even though it is delivered via acentral tube (Song et al., 2008)[13].

This module design included five rectangular plastic-based flatcomponents—one picture frame containing the hollow fiber membranes, twoface boxes and two face plates. To prevent hot brine leakage, there arerectangular gaskets on each side of the membrane-containing pictureframe. Leakage is prevented by having many bolts tightening the wholeassembly together. In the case of two such picture frames placed back toback, the number of face boxes and face plates per picture frame can bereduced to one each.

When such an assembly of two DCMD modules back-to-back (called asingle-pair unit configuration) was used in the pilot plant studies(Song et al., 2008) [13], a single module had ˜0.67 m² membrane surfacearea (based on fiber ID). The total membrane surface area with two suchmodules back-to-back in the assembly was ˜1.34 m². The overall assemblydimensions containing two back-to-back modules were 43 cm×16.5 cm×17.5cm occupying a volume of 12410 cm³. In this configuration, the membranesurface area per unit equipment volume based on the fiber OD (630 μm)instead of the fiber ID (330 μm) is 394 m²/m³.

Leak-free operation of an assembly of such modules is demanding. Thewasted volume in a module is high. Inside each picture frame, half ofthe volume is wasted since the hollow fibers cannot be potted over thewhole thickness of the picture frame. Scale up to larger dimensions isproblematic. In effect, a very large number of such small units need tobe assembled for scale up. Even though it is a hollow fiber-based unit,the effective membrane surface packing density is quite low, resultingin a large footprint and weight in larger-scale plants. Further, thecost naturally goes up in a plate and frame configuration due to so manyflat plates having well-machined surfaces to provide leak-proofoperation, appropriate flow distribution, etc.

SUMMARY

In accordance with embodiments of the present disclosure, exemplaryembodiments are generally directed to a cylindrical cross-flow hollowfiber-based module.

A cylindrical cross-flow hollow fiber-based module configuration canlead to potentially much smoother scale up. A DCMD process is disclosedfor recovering water vapor from brine using a novel cross-flow hollowfiber membrane module in a cylindrical geometry over the temperaturerange of approximately 40 to 95° C. This temperature range is onlyexemplary. It will be understood that other temperature ranges could beemployed. Pressure can also be higher than atmospheric. The porousfluorosiloxane-coated PP hollow fiber membranes in the module weresimilar to those employed in earlier studies. The PP hollow fibers maybe replaced by those of polyvinylidene fluoride (PVDF), and otherhydrophobic polymers such as polytetrafluoroethylene (PTFE),poly(4-methyl-1-pentene (PMP), etc.

Thermally driven membrane distillation-based desalination is becomingattractive especially for concentrated saline waters having scalingsalts. Although a rectangular module with crossflow of hot brine overhydrophobic porous hollow fibers of PP having a porous fluorosiloxanecoating demonstrated extraordinary DCMD performance and resistance tofouling by scaling precipitates, it had a low surface area per unitvolume and posed scale-up problems. A cylindrical hollow fiber devicehas been developed having a surface area per unit volume of 1526 m²/m³which is about four times that of the rectangular modules developedearlier. This surface area per unit volume is only exemplary. It will beunderstood that other surface areas per unit volume could be used. Itcan be scaled up very easily to larger diameters to accommodate largevalues of membrane surface area.

The module has been operated with the hot brine coming into the shellside through a central feed tube either from one end (dead-endconfiguration) or from both ends (split-flow configuration) and goingradially out. The results of numerical simulations of a model developedfor the dead-end configuration provides a reasonable description of theobserved water vapor flux variation with hot brine inlet temperaturewhen the module is operated in the dead-end operational mode. Thesplit-flow configuration could provide a slightly higher flux. In oneembodiment, the hollow fiber length of 45.7 cm in the largest module isalmost twice that of the length of the largest rectangular modulestudied earlier. Simulations of the model show that the membrane watervapor flux performance is in line with what was observed with thelargest rectangular modules studied earlier in a pilot plant. One canincrease the length of the hollow fibers to a few times that of 45.7 cmlength. Further, the internal diameter of the hollow fiber may beincreased appropriately to reduce the distillate side pressure drop.

Any combination and/or permutation of the embodiments are envisioned.Other objects and features will become apparent from the followingdetailed description considered in conjunction with the accompanyingdrawings. It is to be understood, however, that the drawings aredesigned as an illustration only and not as a definition of the limitsof the present disclosure.

BRIEF DESCRIPTION OF THE DRAWINGS

To assist those of skill in the art in using the disclosed systems andmethods, reference is made to the accompanying Figures, wherein:

FIG. 1 is a schematic view of cross-flow module with flow directions atthe shell side and feed brine entering the central feeder tube from bothsides.

FIG. 2 is a schematic view of cross-flow module with flow directions atthe shell side and feed brine entering the central feeder tube from oneside only.

FIG. 3 is a perspective view of an embodiment of a module.

FIG. 4A is a top plan view of the module of FIG. 3 .

FIG. 4B is a cross-sectional view of the module of FIG. 4A taken alongcut line A-A.

FIG. 5A is a cross-sectional view of an embodiment of a module.

FIG. 5B is an enlarged view of the portion of the module of FIG. 5Adenoted by “B”.

FIG. 6 is a schematic diagram of an experimental setup for DCMD with aheat exchanger (HX) and a membrane module.

FIG. 7 is a chart showing change in pressure drop encountered by shellside brine for different brine flow rates for small modules #1, #2 and#3.

FIG. 8 is a chart showing change in water vapor flux with differentbrine flow rates for 1% NaCl solution at different temperatures in smallmodule #3.

FIG. 9 is a chart showing change in water vapor flux with temperature ofa simulated de-oiled produced water in small module #2 for a simulatedproduced water flow rate of 1,800 ml/min.

FIG. 10 is a chart showing experimental and simulation results of watervapor flux for the large module III.

FIG. 11 is a chart showing experimental and simulation results of watervapor flux for the large module III.

FIG. 12 is a chart showing variation of water vapor flux withtemperature for large module I.

FIG. 13 is a schematic illustration of an arrangement of hollow fibersin the larger DCMD module.

FIG. 14 is a mass and energy balance equation for the length of Δx inthe distillate flow direction.

FIG. 15 is a chart showing experimental and simulation results for largemodule III.

FIG. 16 is a chart showing simulation results of flux vs. fiber lengthfor large module III.

FIG. 17A is a chart showing predicted water vapor flux and tube sidepressure drop vs. shell side inlet temperature at two sets of tube sideflow rates.

FIG. 17B is a chart showing predicted water vapor flux and tube sidepressure drop vs. shell side inlet temperature at various tube side flowrates.

FIG. 17C is a chart showing predicted water vapor flux for differentfiber lengths vs. shell side inlet temperature at various tube side flowrates.

FIG. 17D is a chart showing predicted production rate across differentfiber lengths vs. shell side inlet temperature at various tube side flowrates.

DETAILED DESCRIPTION OF EXEMPLARY EMBODIMENTS

In accordance with embodiments of the present disclosure, exemplaryembodiments are generally directed to a cylindrical cross-flow hollowfiber-based module.

Membrane Modules

A basic design of the cylindrical cross-flow membrane module was basedon an understanding of water vapor transfer rates under particularcross-flow conditions in the larger diameter coated hollow fibers usedin previous studies. The hot brine is in a radially outward flowconfiguration through the hollow fiber bed from a central inlet feedertube. The basic module design is schematically shown in FIG. 1 and FIG.2 . Example modules were produced with membrane surface areas of about0.15 m² and 0.6 m². It will be understood that other membrane surfaceareas could be used.

The brine may be introduced into such a module from both ends of thecentral feeder tube as shown in FIG. 1 (Split-Flow Mode). In such acase, the diameters of the holes on the wall of the central feeder tube(CFT) increase from both ends toward the middle where the increase inthe diameter of the hole stops. On the other hand, if the brine isintroduced from one end only as in FIG. 2 (Dead-End Mode), the diametersof the holes could continue to increase from the brine inlet to theother end.

FIG. 3 illustrates an example membrane distillation module 10 inaccordance with the present invention. With reference to FIGS. 3-5B,Module 10 includes a housing 12 with a substantially cylindrical wall 14and first and second ends 16, 18 together defining a chamber 20 with anaxis 22 passing through first and second ends 16, 18. A delivery conduit24, otherwise referred to herein as a “central inlet feeder tube”,extends into chamber 20. In some embodiments, delivery conduit 24extends through chamber 20 either along or substantially parallel toaxis 22. Delivery conduit 24 includes a lumen 26 for conveying fluidinto chamber 20. An inlet 28 opens to lumen 26 of delivery conduit 24.In the illustrated embodiment, inlet 28 may be fluidically coupled to afirst fluid intake pipe 30 for receiving intake fluid flow 32. In someembodiments, intake fluid flow 32 is the feed flow, which, in the caseof DCMD of brine, is heated brine feed flow.

In some embodiments, delivery conduit 24 may extend axially throughchamber 20 to define first and second axially opposed ends 25 a, 25 b ofdelivery conduit 24. Second end 25 b of delivery conduit 24 may includea second inlet 34 opening to lumen 26 of delivery conduit 24, andadjacent to second end 25 b of delivery conduit 24. First inlet 28 maybe adjacent to first end 25 a of delivery conduit 24, such that each ofthe first and second inlets 28, 34 to lumen 26 are external to chamber20. Such an arrangement provides for “spilt-flow of the intake fluidflow, wherein a second fluid intake pipe 36 is fluidically coupled tosecond inlet 34 to convey intake fluid flow 32 b to inlet 34 into lumen26 of delivery conduit 24. In other embodiments, such as thatillustrated in FIG. 1 , delivery conduit 24 may include a closed endaxially opposite from first end 25 a delivery conduit 24 to provide the“dead end flow” described herein and in reference to the schematicillustration of FIG. 2 . Delivery conduit 24 includes an outlet 38opening to lumen 26 of delivery conduit 24, and positioned in chamber20. Outlet 38 permits fluid flow, such as the hot brine flow, out fromlumen 26 into chamber 20. Outlet 38 may be formed in a variety ofconfigurations and mechanisms. In some embodiments, outlet 38 mayinclude one or more apertures 38 a. An aspect of the present inventionis the provision of outlet 38 with a variable fluid flow-throughrestriction in order to more evenly distribute outlet fluid flow fromlumen 26 throughout a length of chamber 20. Without a variable fluidflow-through restriction, pressure drop of the fluid flow through lumen26 would result in unequal fluid outflow rates and poor axial outflowdistribution. To counteract such effects from pressure drop throughlumen 26, delivery conduit 24 may be provided with an outlet 38 thatplaces less fluid flow-through restriction at locations where reducedfluid pressure regimes in lumen 26 are likely to exist, such as withincreased distance from the one or more inlets to lumen 26.Consequently, outlet 38 may have decreasing fluid flow-throughresistance with increasing axial distance from the fluid flow inlet tolumen 26. In the case of the dead end flow regime schematicallyillustrated in FIG. 1 , outlet 38 may provide a range of decreasingfluid flow-through resistance with increasing axial distance along lumen26 from the single fluid flow inlet to lumen 26. In the embodimentillustrated in FIG. 4B, by contrast, for a split-flow regime utilizingfirst and second fluid inlets 28, 34, fluid flow-through resistance foroutlet 38 may decrease with increasing distance from both of inlets 28,34, such as toward a midpoint 40 of delivery tube 24. For the purposeshereof, the midpoint 40 of delivery tube 24 may be the location at whichintake fluid pressure drop from each of the first and second inlets 28,34 is equivalent. In some embodiments, such midpoint 40 may beequidistant from certain structures of module 10. In the illustratedembodiment, outlet 38 includes a plurality of apertures increasing insize toward midpoint 40 of delivery tube 24. Apertures 38 a may assume avariety of configurations to accomplish the variable fluid flow-throughrestriction described above. It is contemplated that other approachesfor creating a variable fluid flow-through restriction for outlet 38 maybe employed, such as density of apertures, valved apertures, and so on.

Module 12 includes a membrane 50 formed by a plurality of hollow fibers52 in chamber 20. For the DCMD of brine, applicant has found that suchhollow fibers 52 are preferably porous and hydrophobic, as described ingreater detail herein. The respective lumens of such hollow fibers 52form a tube side 21 a of chamber 20. A chamber volume external todelivery conduit 24 and hollow fibers 52 in chamber 20 forms a shellside 21 b of chamber 20.

Hollow fibers 52 may be bundled or unbundled and aligned substantiallyaxially in chamber 20. Hollow fibers 52 may substantially circumaxiallysurround delivery conduit 24, or may be arranged in other patternssuitable for the intended DCMD application. In one aspect of theinvention, hollow fibers 52 may be somewhat loosely packed in chamber20, with a packing fraction of less than about 0.5, more preferably lessthan about 0.3, and still more preferably less than about 0.25. In someembodiments, the packing fraction of hollow fibers 52 in chamber 20 isbetween about 0.03 and about 0.25. The relatively loose packing fractionof hollow fibers 52 in chamber 20 permits radial cross-flow of brinefrom delivery conduit 24 over and between individual fibers to maximizeavailable contact surface area of the hollow fiber outer surface to thebrine. The surface area of the plurality of hollow fibers 52 per unitvolume in chamber 20 may be at least about 400 m²/m³.

The relatively loose packing fraction of hollow fibers 52 in chamber 20may also benefit module 10 in reducing a tendency for precipitatebuildup in and around the fiber bundle. DCMD of brine inevitably resultsin some minerals and salts precipitating from the brine. The solidprecipitate can collect on surfaces in chamber 20, particularly wherelow shell side flow is available to sweep the minerals out from thechamber 20. An example low flow area is within the hollow fiber bundle,external to the fibers. Precipitates can deposit on the outer surface ofthe fibers and form barriers to hot brine flow. The reduced brine flowin that area can lead to further precipitate deposition. This cycle ofprecipitate buildup may be diminished with the presently contemplatedlow fiber packing fraction, wherein sufficient void space permitsmovement of the fibers within the bundle. Precipitates are thereforemore likely to be swept out of chamber 20 with the hot brine, instead ofdepositing on the outer surface of the fibers. A particular concern withprecipitate buildup is with salts having relatively sharp crystallinestructures that can damage the hollow fibers. It is therefore beneficialto provide a module construct that promotes removal of precipitatesduring operation.

Hollow fibers 52 may be secured in chamber 20 by epoxy potting. Asillustrated in FIG. 5B, first ends of hollow fibers 52 are secured in afirst potting 54, which seals against an end cap assembly surface 62 toform a first sealed end 56 to shell side 21 b of chamber 20. Hollowfibers 52 are preferably secured by and in first potting 54 to extend atleast partially therethrough, so that the respective lumens forming tubeside 21 a of chamber 20 are open to a first plenum 72 establishedbetween first potting 54 and a first tube side port 74 in housing 12. Byaccess to the lumens of hollow fibers 52, therefore, first plenum 72 isin fluid communication with tube side 21 a of chamber 20 through firstpotting 54.

A similar arrangement may be provided at axially opposite second ends ofhollow fibers 54, wherein the second ends of hollow fibers 52 aresecured in a second potting 58, which seals against an end cap assemblysurface 64 to form a second sealed end 59 to shell side 21 b of chamber20. The lumens forming tube side 21 a of chamber 20 are preferably opento a second plenum 76 established between second potting 58 and a secondtube side port 78 in housing 12. By access to the lumens of hollowfibers 52, therefore, second plenum 76 is in fluid communication withtube side 21 a of chamber 20 through second potting 58.

In preferred embodiments, cold distillate may be fed into tube side 21 aof chamber 20 through one of first and second tube side ports 74, 78.For example, cold distillate may be fed through second tube side port 78as distillate inlet flow 33 a, into second plenum 76, and further intothe respective hollow fiber lumens forming tube side 21 a of chamber 20.The cold distillate is conveyed within the hollow fiber lumens throughsecond potting 58 into chamber 20. The DCMD process permits collectionof condensed water vapor passing through membrane 50 into tube side 21b. Continued flow of the cold distillate (and collected condensed watervapor) is then conveyed through first potting 54 into first plenum 72,and ultimately out from module 10 through first tube side port 74 asdistillate outlet flow 33 b.

Hot brine, as described above, may be fed through one or more of firstand second fluid intake pipes 30, 36, which are fluidically coupled torespective first and second inlets 28, 34 to lumen 26 of deliveryconduit 24. The hot brine may be conveyed in lumen 26 to outlet 38,wherein the hot brine flows radially out from delivery conduit 24 intoshell side 21 b of chamber 20 for cross flow contact with the hollowfibers 52. Flow of the concentrated hot brine exits from shell side 21 bof chamber 20 through one or more shell side ports 92, 94 in housing 12as shell outlet flow 95 a, 95 b.

It should be understood that “brine” is an example feed fluid fortreatment by the membrane distillation module of the present invention.The term “brine” is intended to mean salinated water. Other feed fluids,however, are contemplated as being useful in the present invention.Moreover, the term “distillate” is meant to include any fluid orenvironment useful in driving distillation transport of vapor acrossmembrane 50. Another pertinent term for such fluid may be “condensingfluid”.

Preferably, shell side ports 92, 94 are disposed between first andsecond sealed ends 56, 59 of chamber 20. Moreover, shell side ports 92,94 may be positioned distal from outlet 38 to ensure sufficient brineresidence time in chamber 20 and contact between the hot brine in shellside 21 b and hollow fibers 52 carrying cold distillate in tube side 21a. By positioning shell side ports 92, 94 between first and secondsealed ends 56, 59 of chamber 20, shell side fluid need not be passedthrough at least one of first and second pottings 54, 58. In many pastdesigns, shell side fluid was routed through narrow and even tortuouspassageways in the epoxy potting. Precipitates from the shell side fluidcould accumulate in passageways, leading to restricted flow and evenblockages.

The illustrated embodiment of module 10 includes first and second endcap assemblies 80, 82 at respective ends 16, 18 of housing 12. Each endcap assembly, in the illustrated embodiment, includes a connectioncollar 84 secured to cylindrical wall 14, and an end cap 86 secured toconnection collar 84. A gasket 88 may be disposed in a gasket groove 90to establish a seal between end cap 86 and collar 84. However, it iscontemplated that end cap 86 may be removably secured to connectioncollar 84 for ease of service to module 10.

Connection collar 84 includes an annular recess 85 that forms a locationfor securely receiving a respective one of first and second pottings 54,58. In particular, annular recess 85 forms a shoulder 87 against whichthe potting may be molded. After placement at annular recess 85, thepotting forms a clean transition with the remainder of the boundingsurface of chamber 20, with minimal ledges or pockets in whichprecipitates could gather and accumulate. By positioning shell sideports 92, 94 between first and second sealed ends 56, 59 of chamber 20in the present arrangement, precipitates are likely to be swept out fromchamber 20 through shell side ports 92, 94 without having an opportunityto accumulate at low flow areas of the module.

A system of a plurality of modules may be constructed in series so thatthe hot brine passes directly from a first distillation chamber to asecond distillation chamber, with the shell side outlet of a firstmodule being fluidically coupled to a delivery conduit in a secondmodule. A number of modules may therefore be placed in series toincrease total concentration of the brine, and recovery of water vaporfrom the original feed stream.

In one embodiment, the hollow fibers are porous hydrophobicpolypropylene (PP) of internal diameter (ID) 330 μm, wall thickness 150μm having a pore size of 0.6 μm and a porosity of 0.6+(Membrana,Charlotte, N.C.). On the outside surface of these hollow fibers there isa light plasma-polymerized fluorosiloxane coating having pores somewhatlarger than those of the PP substrate. Porous hydrophobic hollow fibersof any suitable material may also be used with appropriateplasma-polymerized fluorosiloxane coating on the outside surface. Thesehollow fibers may have other dimensions of their IDs and wallthicknesses as well.

Initially a few small modules (#1, #2, #3) were fabricated using poroushydrophobic polyvinylidene fluoride (PVDF) hollow fibers (Arkema Inc.,King of Prussia, Pa.) as a substrate instead of the coated PP hollowfibers. The performance of these modules guided the development of thedesign for larger modules. All PVDF hollow fiber-based modules weretested in a low temperature DCMD set up (Li and Sirkar, 2004 [11]) andsubsequently, the changes needed were made to improve the DCMDperformance.

The fiber length in both types of larger modules was 45.7 cm. It will beunderstood that the length of the fibers may vary. The module havinglower membrane surface area included fewer layers of hollow fiberswrapped around a central hot brine inlet tube having holes drilled ontheir surface for introducing the hot brine feed into the shell sidearound the hollow fibers. The hole diameters may increase with distancefrom the tube inlet(s). The module having a larger membrane surface areahas a deeper layer of hollow fibers in the radial brine flow directionto achieve˜4 times higher membrane surface area. Details of these hollowfiber membrane modules are provided in Table 1.

TABLE 1 Details of different membrane modules and hollow fibers SmallSmall Small Large Large Large Large module module module module modulemodule Module Particulars #1 #2 #3 I II III IV Membrane PVDF PVDF PVDFCoated Coated Coated PVDF type PP* PP* PP* Fiber ID (μm) 692 692 692 330330 330 692 Fiber OD 925 925 925 630 630 630 925 (μm) No. of fibers 1515 15 316 316 1266 600 Effective 15.5 15.5 15.5 45.7 45.7 45.7 45.7fiber length (cm) Effective 50.4 50.4 50.4 1500 1500 6000 5950 membranesurface area (cm²)** Fiber packing N/A 0.03 0.03 0.13 0.13 0.24 0.25fraction Fiber surface N/A N/A N/A 1120 1120 1526 1526 area per unitvolume (m²/m³)*** CFT ID 0.5 0.5 0.5 1.3 1.3 2.5 2.5 (cm)**** Module IDN/A 2.0 2.0 2.8 2.8 5.2 5.2 (cm) Module 15.5 15.5 15.5 45.7 45.7 45.745.7 length (cm) *Fluorosiloxane coated; **Based on fiber innerdiameter; ***Based on fiber outer diameter; ****CFT—Central Feeder Tube

The shell of the tested DCMD modules was fabricated from standard sizeschedule 40 PVC pipe. The end caps for both of the larger size moduleswere essentially identical. Standard PVC fittings selected were PVCcemented to the inlet and outlet pipes. The goal was to make a sturdy,light, and inexpensive module that is easy to handle and has much moremembrane surface area per unit volume. Further, the module should becapable of being connected easily to other modules. It should be notedthat there are no bolts to secure the end caps to the body.

Other design considerations include essentially no shoulder at theshell-side brine outlet locations at every shell-side outlet. Suchshoulders are locations where precipitates of scaling minerals couldaccumulate leading to a buildup which could even extend to the outermostlayers of the hollow fibers. These sections of hollow fibers areunlikely to be subjected to the beneficial effects of crossflow andcould therefore be potentially vulnerable to wetting-induced brineleakage (see the results shown in FIG. 10 of He et al. (2009a) [16]).

Some design items required experimental verification of their utility.One such item is the shell-side pressure drop as brine flows radiallyoutward from the central tube. It depends on the number of holes and thesize of the holes on the periphery of the central tube among others.That is why smaller radial cross-flow modules were fabricated and testedfor their DCMD performance. Additional design items involve the packingdensity of fibers, their possibility of oscillation at higher radialflow velocities and the gap at the outer periphery between the fiberbundle and the shell ID. In one embodiment, the module design allowedmembrane surface area packing density upwards of 1,500 m²/m³.

Considerations on Membrane Module Design

Smaller Membrane Modules

In each of the three small membrane modules #1, #2 and #3, the sectionhaving perforated length at the center of the central tube of diameter0.5 cm was 15 cm long. The central tube of these modules was made withperfluoroalkoxyethylene (PFA) tubing. In module #1, holes of only onesize were introduced. In modules #2 and #3, two different hole sizeswere created with larger holes in the middle part of the perforatedlength and the smaller holes on each side of the larger hole region. Thehole sizes in module #3 were larger than those in module #2 whose holesizes were larger than those of module #1. The goal was to study how toreduce shell-side brine pressure drop. The number of PVDF hollow fibersin each module was 15, providing an ID-based effective surface area of50.47 cm² and an effective length of 15.6 cm. The hollow fiberdimensions and other information are provided in Table 1. Each modulewas designed so that hot feed solution could enter from both ends of themodule (split-flow).

Larger Membrane Modules

The two larger membrane modules I and II were essentially identical. Inlarger modules I and II, the section having perforated length at thecenter of the 1.27 cm diameter central tube was 45.7 cm long. Thelargest module III with 4 times larger membrane surface area than thatof the larger modules I and II was fabricated using standard sizeschedule 40 PVC pipe and had a 2.54 cm OD perforated central tube in a5.23 cm diameter PVC pipe acting as shell. This module contained 1266fluorosiloxane coated porous PP hollow fibers with an effective lengthof 45.7 cm (18 in) and effective ID-based membrane surface area of 0.6m².

Experimental Details for DCMD Studies

Two experimental setups were used for finding out the DCMD performancesof various modules. The smaller experimental setup shown in FIG. 6 wasemployed and described in Lee et al. (2011) [5]. The only difference isthat only one cylindrical membrane module was used and not acountercurrent cascade of four small rectangular cross flow modules. Thelarger experimental setup has been described in detail in Song et al.(2007) [12]. The water vapor flux is defined as

$\begin{matrix}{{{Water}{vapor}{flux}\left( \frac{kg}{m^{2} - {hr}} \right)} = \frac{{Water}{vapor}{collected}({kg})}{{Membrane}{area}{based}{on}{ID}\left( m^{2} \right) \times {time}({hr})}} & (1)\end{matrix}$

The conductivity on the distillate side was measured using aconductivity meter (Orion 115 A

, ThermoElectron, Waltham, Mass.). All experiments were performed withhot brine of 1 wt % NaCl. A few experiments were performed with a smallmodule using a simulated produced water (synthetic water simulating thecomposition of the Post WEMCO stream (total dissolved solids, 7622 mg/L)(Singh et al. [21])).

Results and Discussion

DCMD Performances of Small Membrane Modules

Experiments were performed with feed brine entering from one end or bothends of the module keeping it in horizontal as well as in verticalposition. In all experiments, hot brine was passed through the centraltube for radial emission across the porous hollow fibers having cold DIwater flowing through them. Due to the very small size of holes in thecentral feed tube, pressure drop encountered by the shell-side brine wassignificant. In module #1, the pressure drop increased from 55.12 kPa (8psi) to 110.2 kPa (16 psi) as the flow rate was increased from 0.8 L/minto 1.5 L/min. Due to the modification in the central tube design,pressure drop encountered by the shell-side brine in module #3 was muchlower compared to those in small module #1 and small module #2; it wentup from around 0 kPa to 41.34 kPa (6 psi) as the brine flow rate wasincreased from 0.8 L/min to 1.8 L/min. These results are shown in FIG. 7. Since high pressure drop is not desirable for an energy-efficientprocess, further changes were introduced in the design of the largermodules including a larger diameter of the central tube itself.

The DCMD performance of the small module #3 was studied for differentbrine temperatures and different brine flow rates with 1% NaCl feedsolution. As shown in FIG. 8 , it was found that water vapor flux was ashigh as 9.9 kg/m²-hr for brine feed at 90° C. and 1.8 L/min. Water vaporflux increased from 3.1 kg/m²-hr to 5.9 kg/m²-hr as the brine flow ratewas increased from 0.8 L/min to 1.8 L/min for the brine feed at 85° C.Similarly, for a brine temperature of 90° C., water vapor flux increasedfrom 5.3 kg/m²-hr to 9.9 kg/m²-hr as the brine flow rate was increasedfrom 0.8 L/min to 1.8 L/min. The conductivity on the distillate side wasconstant for all experiments indicating that there was no salt leakagefrom the brine side to the distillate side through the membrane.

The performance of the small module #2 was studied with simulatedde-oiled produced water at different temperatures. Water vapor fluxincreased from 4.2 kg/m²-hr to 6.6 kg/m²-hr as the feed temperature wasincreased from 85° C. to 91° C. (FIG. 9 ). Water vapor fluxes obtainedwith the simulated produced water were similar to those obtained with 1%NaCl solution. However, for the simulated produced water, theconductivity on the distillate side increased with time for differentfeed temperatures. After three hours, the conductivity became constantwith time irrespective of the feed water temperature. As observed inearlier studies with Chevron-supplied produced water (Singh et al.[21]), the distillate side conductivity increase had very little to dowith salt leakage or pore wetting. It was primarily due to dissolved CO₂coming to the distillate water from bicarbonates in the feed solutiondissociating at the higher temperatures as has been already described inSingh et al. [21] and He et al. [16].

DCMD Performances of Larger Membrane Modules

It is useful to note at the beginning the values of the membrane surfacearea per unit volume for the three larger membrane modules based on thefiber outside diameter. As Table 1 shows, these are 1120, 1120, 1526m²/m³ for Modules I, II and III, respectively. These are a few timeslarger than that of the rectangular modules used in earlier pilot plantstudies (Song et al., 2008) [13]; the surface area for module III isalmost 4 times larger. Two module configurations were tested in so faras brine introduction is concerned. In Dead-End Mode, hot brine is fedthrough the bore of the 1.27 cm diameter central feed distribution tube,and is emitted radially through the holes in the wall to flow radiallyacross the porous hollow fibers and out from the shell side. The otherend of the central distribution tube is closed. In Spilt-Flow Mode, hotbrine is introduced from both ends of the central distribution tube.

In FIG. 10 , the experimental flux results for the large module III areshown as a function of brine inlet temperature varying from 75.2 to84.4° C. in Dead-End Mode of operation; the brine flow rate was 18L/min, and the inlet brine temperature was varied between 75.2-84.4° C.Distillate flowed concurrently with brine flow direction in the CFT. Thedistillate was fed at 23.9-24.8° C. and 2.5 L/min. FIG. 11 illustratesthe experimental results for the same system for a lower brine flow rateof 15 L/min and a lower brine inlet temperature range of 59.9-79.9° C.The water vapor fluxes are slightly lower due to lower brine velocityand lower brine inlet temperature. These data were taken in the largerexperimental setup. Distillate side conductivity was constant indicatingthat there was no salt leakage from the brine side to the distillateside through the hollow fiber pores.

FIG. 12 illustrates experimental data of the large module I inSplit-Flow Mode at a brine flow rate of 18 L/min in the range of 80 to91° C. for brine-in temperature. Distillate-in temperature wasmaintained at ˜22° C. for a flow rate of 0.9 L/min. Water vapor fluxincreased from 12.5 kg/m²-hr to 16.8 kg/m²-hr as the brine-intemperature was increased from 80.5° C. to 91° C. Brine-in pressure wasaround 27.56 kPag (4 psig) during all experiments at the flow rate of 18L/min.

Modeling of DCMD Performances of Large Membrane Modules

The performance modeling is focused on a radial cross flow hollow fibermembrane module. Sengupta et al. (1998) [22] modeled degassing of waterflowing radially and counter-currently on the shell-side due to thepresence of a baffle; there was no modeling involved on the permeateside. Appropriate equations have been developed for a mathematical modelof direct contact membrane distillation with the hot brine entering inthe dead end mode. First consider the pattern of hollow fibers incircles around the central core tube (shaded) bringing in the hot brine(FIG. 13 ) which spreads out radially throughout the fiber bundle. Inany such fiber bundle, focus on one hollow fiber in one particularlayer. The various terms needed for mass and energy balances around onehollow fiber at a local section at a distance of x from the distillateentry point are given below (see notation).

Consider now a differential slice of the DCMD module with radius r_(j)and radial width dr_(j) identified as the j^(th) fiber layer. The areaof this annulus is approximately 2πr_(j) dr_(j). The number of hollowfiber dn_(j) inside this slice is obtained from relations (2a) and (2b)given below:

$\begin{matrix}{{dn}_{j} = {\frac{f_{p}\left( {2\pi r_{j}{dr}_{j}} \right)}{\frac{\pi}{4}d_{o}^{2}} = \frac{8f_{p}r_{j}{dr}_{j}}{d_{o}^{2}}}} & \left( {2a} \right)\end{matrix}$ $\begin{matrix}{n_{j} = {\frac{4f_{p}}{d_{o}^{2}}r_{j}^{2}}} & \left( {2b} \right)\end{matrix}$

Therefore, in the circle of radius r_(j) the number of hollow fiberslocated with their center at radius r_(j) is n_(j). Here f_(p) is thefractional packing density of N number of hollow fibers (of diameterd_(o)) in the shell side of diameter d_(s) (around the central core tubeof diameter d_(t)); it is defined as

$\begin{matrix}{f_{p} = {\frac{N\frac{\pi}{4}d_{o}^{2}}{{\frac{\pi}{4}d_{s}^{2}} - {\frac{\pi}{4}d_{t}^{2}}} = \frac{Nd_{o}^{2}}{d_{s}^{2} - d_{t}^{2}}}} & (3)\end{matrix}$

As the value of r_(j) increases, the number of fibers in that layerincreases with the square of the radius of the radial location.

Mass Balance on Jth Layer with n_(j) Number of Hollow Fibers

The difference in the distillate mass flow rate in the j^(th) layer offibers is equal to the difference in brine mass flow rate over thej^(th) layer of hollow fibers (see FIG. 14 ) (Note: Distillate flow ishere co-current with brine flow direction in CFT):

$\begin{matrix}{{\int_{0}^{L}{\left\lbrack {N_{v,j}(x)} \right\rbrack n_{j}\pi d_{\ln}{dx}}} = \text{?}} & (4)\end{matrix}$ $\begin{matrix}\left. \left. {\text{?},{{in} - F_{f,{out}}}} \right) \right|_{j} & (5)\end{matrix}$ ?indicates text missing or illegible when filed

Here d_(In) is defined as:

$\begin{matrix}{d_{in} = \frac{d_{o} - d_{i}}{\ln\left( \frac{d_{o}}{d_{i}} \right)}} & (6)\end{matrix}$

Further N_(v,j)(x) is the water vapor mass flux in the j_(th) layer withn_(j) number of hollow fiber at any x and k_(m) is the water vapor masstransfer coefficient through the membrane:

N _(v,j)(x)=k _(m)(P _(fm,j)(x)−P _(pm,j)(x))  (7)

Here the water vapor partial pressures P_(fm,j)(x) and P_(pm,j)(x) canbe expressed using Antoine equation (Smith et al., 2001) [23]:

$\begin{matrix}{{P_{{fm},j}(x)} = {10^{3}{\exp\left( {{1{6.2}60} - \frac{379{9.8}9}{{T_{{fm},j}(x)} + {27{3.1}5} - {4{6.8}}}} \right)}}} & (8)\end{matrix}$ $\begin{matrix}{{P_{{pm},j}(x)} = {10^{3}{\exp\left( {{1{6.2}60} - \frac{379{9.8}9}{{T_{{pm},j}(x)} + {27{3.1}5} - {4{6.8}}}} \right)}}} & (9)\end{matrix}$

Heat Balance on Jth Layer with n_(j) Number of Hollow Fibers

The heat gain rate of distillate is equal to the heat loss rate ofbrine:

$\begin{matrix}{\begin{matrix}\left. {\int_{0}^{x}{{dQ}(x)}} \right|_{j} & {= {C_{p}\left\lbrack i \right.}}\end{matrix}\text{?}} & (10)\end{matrix}$ $\begin{matrix}{\begin{matrix} = & {\frac{1}{x}C_{p}{\int_{0}^{x}\left( \lbrack \right.}}\end{matrix}\text{?}} & (11)\end{matrix}$ ?indicates text missing or illegible when filed

Shell Side Brine Heat Transfer

$\begin{matrix}{\left. \frac{d{Q(x)}}{dx} \right|_{j} = {h_{f,j}n_{j}\pi{d_{o}\left( {{T_{{fo},j}(x)} - {T_{{fm},j}(x)}} \right)}}} & (12)\end{matrix}$

The heat transfer coefficient h_(f,j) in the brine side could beexpressed based on Z̆ukauskas equation (Z̆ukauskas, 1987)[24] for givenvalues of Re_(o) and Pr_(o) (Song et al., 2007)[12]:

$\begin{matrix}{{{Nu_{f,j}} = {\frac{h_{f,j}d_{o}}{k_{o}} = {104{Re}_{o}^{0.4}P{r_{o}^{0.36}\left( \frac{Pr_{o}}{Pr_{w}} \right)}^{{0.2}5}F_{c}\left( {{Re} < 40} \right)}}}} & (13)\end{matrix}$ $\begin{matrix}{{Nu}_{f,j} = {\frac{h_{f,j}d_{o}}{k_{o}} = {0.71{Re}_{o}^{05}P{r_{o}^{{0.3}6}\left( \frac{Pr_{o}}{Pr_{w}} \right)}^{{0.2}5}F_{c}\left( {{Re} > 40} \right)}}} & (14)\end{matrix}$

Where

$\begin{matrix}{{{Re_{o}} = \frac{d_{o}u_{o,j}\rho_{o}}{\mu_{o}}},{{{Pr_{o}} = \frac{C_{po}\mu_{o}}{k_{o}}};{\Pr_{w} = \frac{C_{pw}\mu_{w}}{k_{w}}}}} & \left( {15a} \right)\end{matrix}$ $\begin{matrix}{u_{o,j} = {\frac{F_{f,j}}{3600} \times \frac{1}{{{\pi d}_{j}L} - {n_{j}{\pi d}_{o}L}}}} & \left( {15b} \right)\end{matrix}$

Tube Side Distillate Heat Transfer

$\begin{matrix}{\left. \frac{d{Q(x)}}{dx} \right|_{j} = {h_{p}n_{j}\pi{d_{i}\left( {{T_{{pm},j}(x)} - {T_{{p1},j}(x)}} \right)}}} & (16)\end{matrix}$

The distillate heat transfer coefficient h_(p) is based on theSeider-Tate′ equation (Seider and Tate, 1936) [25]:

$\begin{matrix}{{Nu_{p}} = {\frac{h_{p}d_{i}}{k_{i}} = {186\left( \frac{d_{i}}{L} \right)^{033}\left( {{Re}_{i}Pr_{i}} \right)^{033}\left( \frac{\mu_{i}}{\mu_{wi}} \right)^{0.14}}}} & (17)\end{matrix}$ $\begin{matrix}{{{Re}_{i} = \frac{d_{i}u_{i}\rho_{i}}{\mu_{i}}};{\Pr_{i} = \frac{C_{pi}\mu_{i}}{k_{i}}}} & \left( {18a} \right)\end{matrix}$ $\begin{matrix}{u_{i} = {\frac{F_{d}}{3600} \times \frac{1}{\frac{N\pi}{4}d_{i}^{2}}}} & \left( {18b} \right)\end{matrix}$

Heat Transfer Across the Hollow Fiber Membrane

$\begin{matrix}{\left. \frac{d{Q(x)}}{dx} \right|_{j} = {{h_{m}n_{j}{d_{\ln}\left( {{T_{{fm},j}(x)} - {T_{{pm},j}(x)}} \right)}} + {N_{v,j}n_{j}{{\pi d}_{j}\left( {\Delta{H_{v}\left( {{T_{{pm},j}(x)} + {C_{p,j}{T_{{pm},j}(x)}}} \right)}} \right.}}}} & (19)\end{matrix}$

Where

$\begin{matrix}{N_{v,j} = \frac{\int_{0}^{L}{\left\lbrack {N_{v,j}(x)} \right\rbrack{dx}}}{L}} & (20)\end{matrix}$

From the relations given above, one can get the following:

$\begin{matrix}{\text{?}n_{j}\pi d_{\ln}L} & (21)\end{matrix}$ $\begin{matrix}{\left. T_{f} \middle| {}_{j + 1}(x) \right. = \frac{\left. {m_{j}C_{P}T_{f}} \middle| {}_{j}{(x) - \left( \frac{dQ}{dx} \right)} \right|_{j}}{\left( {m_{j} - {N_{v,j}n_{j}\pi d_{\ln}L}} \right)C_{p}}} & (22)\end{matrix}$ ?indicates text missing or illegible when filed

Given the flow rate and temperature of brine and distillate in thej^(th) layer at any x, the values of T_(fm,j)(x), T_(pm,j) (x),T_(pl,j)(x), (x), and F_(pl,j)(x) can be calculated from the equationsgiven above, along with the boundary condition Q(0)|_(j)=0 using MATLAB.This assumes that the heat transfer coefficients on the brine side andthe distillate side are known. The values of T_(fl,j)(x), Q(x)|_(j),P_(fm,j)(x), P_(pm,j) (x) and F_(bo) (x) can then be solved. A detailednotation section has been provided.

Simulations of the hollow fiber DCMD module performance in rectangularcross-flow were carried out earlier by Song et al. (2008) [13]. Thosesimulations had only one adjustable parameter namely, k_(m), themembrane water vapor mass transfer coefficient; its values are availablein Sirkar and Song (2009) [26]. In the simulations carried out here,k_(m) is also the only adjustable parameter. Table 2 lists the valuesused which are not too far apart from those used by Sirkar and Song(2009)[26]. The modeling used the input values V_(b0), T_(b0), V_(d0),T_(d0), and the details of the module geometry and fiber dimensions andproperties.

TABLE 2 Values of the parameters used in model simulations for Dead-EndMode Reference Temperature T₀ 273.15 K Liquid water heat capacity, C_(p)4.1863 kJ/kg-C Liquid water density 1 g/cm³ Latent heat of evaporation2257 kJ/kg Thermal conductivity for polypropylene, k_(pp) 0.17 W/m-KThermal conductivity for air, k_(air) 0.025 W/m-K Mass transfercoefficient k_(m) for large module I 0.0017 kg/m²/h/Pa Mass transfercoefficient k_(m) for large module III 0.0033 kg/m²/h/Pa

Comparison of Simulation Results with Experimental Results

The model illustrated above was based on the hot brine fed at one end ofthe central tube in the Dead-End Mode. In FIG. 10 , simulation resultsfor the large module III are compared with experimental results of watervapor flux as a function of the brine temperature varying over75.2-84.4° C. in this particular mode of operation i.e., the brine flowrate of 18 L/min in Dead-End Mode. Further, the distillate flowdirection was concurrent with respect to the brine flow direction in theCFT. The distillate coming in at 23.9-24.8° C. had a flow rate of 2.5L/min. The simulation results are somewhat higher but not too far fromthe observed results. FIG. 11 illustrates the corresponding scenario fora lower feed brine flow rate of 15 L/min at a lower brine temperature.Here also the simulation results are somewhat higher than theexperimentally observed values but not too far apart. The value of k_(m)used for the large module III (see Table 2) is close to the value of0.0028 kg/m²/h/Pa used by Sirkar and Song (2009)[26].

The simulation results for the Dead-End Mode were compared in FIG. 15with the data obtained from the large module III in Split-Flow Mode at ashell-side brine flow rate of 18 L/min and a brine inlet temperature of76-80° C. If the performances of the module in two different feed brineflow configurations are compared in one embodiment at around 80° C. fora brine flow rate of 18 L/min, the Split-Flow mode provided a somewhathigher water vapor flux than the Dead-End mode. However, littledifference at lower feed brine temperatures was observed. The Split-Flowmode does provide a more uniform brine flow distribution on the shellside and therefore a better performance.

It is important to note from these figures that the simulation resultsobtained in the dead-end mode are significantly higher than the observedvalues at lower brine temperatures; however at higher temperatures thesimulation results appear to be closer to the experimental values. Thisdeviation is due to a weak temperature dependence of the adjustableparameter k_(m); lower k_(m) values used for lower feed brinetemperatures would bring the simulation results closer to theexperimentally observed values at lower temperatures.

It is useful to explore the effects of the length of the hollow fibersin such a module via simulations in Dead End mode. The large module IIIused here has an effective fiber length of 45.7 cm. It will beunderstood that other fiber lengths could be used. FIG. 16 illustratesthe water vapor flux as a function of the hollow fiber length. Table 3provides numerical values of a variety of relevant quantities. Thesecalculations show that as the hollow fiber length is reduced, watervapor flux is increased considerably while the distillate outlettemperature rise is reduced, contributing to an increase in the flux.For a perspective, the simulation results shown in FIG. 16 can becompared with the performance of rectangular cross-flow modules having alength of 24.1 cm of the hollow fibers used in pilot plant studies (Songet al., 2008[13]). The simulations of FIG. 16 suggest a flux of 24.5kg/m²-h; this value is close to the pilot plant data for the feed brinetemperature range being considered.

TABLE 3 Detailed temperature and flux information for large module IIIsimulations per FIG. 16. Fiber Fiber length length T_(bi) T_(bo) T_(di)T_(do) Flux (cm) (inch) (° C.) (° C.) (° C.) (° C.) (kg/m²-h) 45.7 1879.5 74.0 25.2 32.5 14.9 43.2 17 79.5 74.0 25.2 32.2 15.0 40.6 16 79.574.1 25.2 31.9 15.7 38.1 15 79.5 74.1 25.2 31.5 16.6 35.6 14 79.5 74.225.2 31.1 17.5 33.0 13 79.5 74.3 25.2 30.6 18.7 30.5 12 79.5 74.4 25.229.9 19.9 27.9 11 79.5 74.5 25.2 29.1 21.3 25.4 10 79.5 74.6 25.2 28.322.8 22.9 9 79.5 74.8 25.2 27.2 24.5 20.3 8 79.5 75.0 25.2 25.8 26.5

Simulations for Increased Fiber ID in Cylindrical Cross-Flow FiberModules

Table 1 provides details of two large modules: module III studied so farand a hypothetical one, module IV, where the HFM ID is 692 μm (see PVDFhollow fibers). FIG. 17 a illustrates predicted water vapor flux andtube side pressure drop for HFMs with d_(i)/d_(o) being 330/630 μm.Water vapor flux increased with increasing tube side flow rate. The tubeside pressure drop was 46.2 kPa (6.7 psi) for a distillate flow rate 2.5L/min. It doubled to 92.5 kPa when distillate flow rate was doubled to 5L/min. However, modeling with fibers (Table 1) having d_(i)/d_(o),692/925 μm (FIG. 17 b ) shows much lower pressure drop; the figure alsolists the predicted water vapor flux. NOTE: Tube side pressure drop wasdrastically decreased from 46.2 kPa to 5 kPa at V_(d0)=2.5 L/min; itdecreased further from 92.5 kPa to 10.1 kPa at V_(d0)=5 L/min. Watervapor flux is still comparable to that from the fibers with d_(i)/d_(o)values of 330/630 μm.

Larger tube side flow rates of V_(d0)=10 L/min and V_(d0)=22.5 L/minwere also used for simulation with larger ID fibers (FIG. 17 b ). AtT_(bo)=90° C., water flux was 34.4 kg/m²-h for V_(d0)=10 L/min with tubeside pressure drop 20.2 kPa (2.9 psi); it was 37.6 kg/m²-h forV_(d0)=22.4 L/min with tube side pressure drop 45.5 kPa (6.6 psi). Thisis another advantage of higher tube-side flow rate: a lower distillatetemperature increase results in substantially higher water vapor flux.

The effects of fiber length on simulated water vapor flux for the largerID HFM diameter are shown in FIG. 17 c . At T_(bo)=90° C., as fiberlength was doubled to 91.4 cm, water vapor flux decreased from 19.1kg/m²-h to 10.3 kg/m²-h at V_(d0)=2.5 L/min: tube side pressure drop was10.1 kPa (1.5 psi). At V_(d0)=5 mL/min, the value of water vapor fluxdecreased from 28.2 kg/m²-h and to 17.2 kg/m²-h, while tube sidepressure drop was increased to 20.2 kPa (2.9 psi). The value of watervapor flux decreased from 34.4 kg/m²-h and to 23.4 kg/m²-h as fiberlength was doubled at V_(d0)=10 L/min with a tube side pressure drop of40.4 kPa (5.9 psi). This is attributed to the effectively lowertemperature difference between the two sides of the HFM due to thelonger retention time of distillate stream within the longer HFMs;therefore, the temperature of the distillate increased further. For afiber length of 91.4 cm and flow rate 22.5 L/min, the tube side pressuredrop was as high as 90.0 kPa (13.2 psi). Therefore, water vapor fluxsimulation was not performed.

The simulated results of the effect of fiber length on the waterproduction rate for HFMs with larger d_(i), d_(o) are shown in FIG. 17 d. At T_(b0)=90° C., it is 22.4 L/h for V_(d0)=22.5 L/min, and 20.5 L/hfor V_(d0)=10 L/min for L=45.7 cm; it is 27.9 L/h for V_(d0)=10 L/minfor a fiber length of 91.4 cm, and it was 16.8 L/h for V_(d0)=5 L/min.For a fiber length of 91.4 cm, it is 12.3 L/h for V_(d0)=5 L/min. Animproved estimate of fiber properties and operating conditions are:d_(i), 692 μm; d_(o), 925 μm; N, 600; L, 94.1 cm; V_(b0), 22.5 L/min;V_(d0)=10 L/min; T_(b0), 90° C.; Tao, 20° C., which will give watervapor flux of 23.4 kg/m²-h, water production rate of 27.9 L/h and apressure drop of 40.4 kPa (5.9 psi). Compared to the experimentaloperation conditions (pressure drop 46.2 kPa, 6.7 psi), the waterproduction rate will be 2.3 times higher.

Advantages of Cylindrical Cross-Flow Hollow Fiber Modules in DCMD

The HFM surface area packed in this new, compact and light-weightcylindrical module has a reasonable value of 1526 m²/m³ based on fiberOD; it is 4-5 times larger than that in the rectangular module dependingon estimation based on the fiber OD or fiber ID. The cylindrical modulecan be easily scaled up to 10-20 cm shell diameter and accommodate a fewtimes to more than an order of magnitude higher membrane surface area.Putting a large number of such modules together in a countercurrentcascade and for larger production rates should be straightforward. Theshell-side design automatically sweeps away scaling salt precipitates.The best features of the rectangular cross-flow HFM modules have beenretained; their cumbersome and costly design features inhibitingscale-up for higher production levels have been eliminated. Compared to18 bolts and nuts used in each module used for pilot plant studies (Songet al. [13]), the cylindrical modules need just a few pipe fittings anda few Phillips screws allowing rapid assembly.

There is an additional specific advantage of the Split-Flow Mode ofoperation when the cylindrical cross-flow modules are coupled togetherin a countercurrent cascade (Lee et al., 2011) [5]. The cooled brineexiting through two shell-side outlets of a module operating at a highertemperature in the cascade can easily enter the next module operating ata lower temperature in the cascade from the two sides of the centralfeeding tube.

Thermally driven membrane distillation-based desalination is becomingattractive especially for concentrated saline waters having scalingsalts. Although a rectangular module with crossflow of hot brine overhydrophobic porous hollow fibers of PP having a porous fluorosiloxanecoating demonstrated extraordinary DCMD performance and resistance tofouling by scaling precipitates, it had a low surface area per unitvolume and posed scale-up problems. A cylindrical hollow fiber devicehas been developed having a surface area per unit volume of 1526 m²/m³which is about four times that of the rectangular modules developedearlier. It can be scaled up easily to larger diameters to largediameters and high membrane surface areas.

The module has been operated with the hot brine coming into the shellside through a central feed tube either from one end (dead-end mode) orfrom both ends (split-flow mode) and going radially out. The results ofnumerical simulations of a model developed for the dead-end mode providea reasonable description of the observed water vapor flux variation withhot brine inlet temperature when the module is operated in the dead-endoperational mode. The split-flow mode provides a slightly higher flux.The hollow fiber length of 45.7 cm in the largest module in oneembodiment is almost twice that of the length of the largest rectangularmodule studied earlier. Simulations of the model show that the membranewater vapor flux performance in embodiments is in line with what wasobserved with the largest rectangular modules studied earlier in a pilotplant. Simulations further show that larger ID hollow fibers will reducethe distillate-side pressure drop drastically, accommodate a much higherdistillate flow rate leading to higher fluxes and a higher waterproduction rate per module with longer HFMs.

While exemplary embodiments have been described herein, it is expresslynoted that these embodiments should not be construed as limiting, butrather that additions and modifications to what is expressly describedherein also are included within the scope of the invention. Moreover, itis to be understood that the features of the various embodimentsdescribed herein are not mutually exclusive and can exist in variouscombinations and permutations, even if such combinations or permutationsare not made express herein, without departing from the spirit and scopeof the invention.

1. A method for desalination of liquid brine, comprising: (a) providinga membrane distillation module having: (i) a housing with asubstantially cylindrical wall and first and second ends togetherdefining a chamber with an axis passing through the first and secondends; (ii) a delivery conduit extending into said chamber, and having afirst inlet opening to a lumen of said delivery conduit, the first inletpositioned external to said chamber, and an outlet opening to the lumenand positioned in said chamber; (iii) a membrane formed by a pluralityof unbundled porous hydrophobic hollow fibers in said chamber, lumens ofsaid hollow fibers forming a tube side of the chamber; (iv) a firstpotting for securing first ends of said hollow fibers, and forming afirst sealed end to the shell side of said chamber; (v) a second pottingfor securing second ends of said hollow fibers, and forming a secondsealed end to the shell side of said chamber; a chamber volume externalto said hollow fibers and said delivery conduit and between the firstand second sealed ends of the chamber forming a shell side of thechamber in which shell side fluid is contained by and between the firstand second sealed ends; (vi) a shell side port in said housing openingto the shell side of said chamber, said shell side port disposed betweenthe first and second sealed ends; (vii) a first tube side port in saidhousing opening to a first plenum that is in fluid communication withthe tube side of said chamber through said first potting; and (viii) asecond tube side port in said housing opening to a second plenum that isin fluid communication with the tube side of said chamber through saidsecond potting; (b) flowing the liquid brine at a first temperaturethrough the shell side of the chamber; and (c) flowing a second liquidat a second temperature through the tube side of the chamber, whereinthe first temperature is higher than the second temperature.
 2. Themethod as in claim 1, including flowing the second liquid to the tubeside of the chamber through one of said first and second tube sideports.
 3. The method as in claim 1, including flowing the liquid brineto the shell side of the chamber through the outlet of the deliveryconduit.
 4. The method as in claim 3 wherein said membrane separates thetube side of the chamber from the shell side of the chamber.
 5. Themethod as in claim 4 wherein the liquid brine passes over the hollowfibers in radial cross flow to pass desalinated water vapor throughmembrane pores of the membrane to the tube side of the chamber.
 6. Themethod as in claim 1 wherein said delivery conduit extends axiallythrough said chamber to define first and second axially opposed ends ofthe delivery conduit, and includes a second inlet opening to thedelivery conduit lumen positioned external to said chamber and adjacentto the second end of the delivery conduit, wherein the first inlet ispositioned adjacent to the first end of the delivery conduit.
 7. Themethod as in claim 6, including flowing the brine into both of the firstand second inlets of the delivery conduit, and through the outlet of thedelivery conduit.
 8. The method as in claim 7 wherein said deliveryconduit extends axially through said first and second ends of saidhousing.
 9. The method as in claim 1 wherein said delivery conduitincludes a closed end axially opposite from a first end of the deliveryconduit, adjacent to which first end is the inlet opening to the lumenof the delivery conduit.
 10. The method as in claim 1 wherein said firstand second ends of said housing include respective end caps that areremovably engagable with said substantially cylindrical wall.
 11. Themethod as in claim 10 wherein said end caps are removably engagable withsaid substantially cylindrical wall without separate fasteners.
 12. Themethod as in claim 1 wherein a packing fraction of the plurality ofporous hydrophobic hollow fibers in said chamber is less than 0.25.